U.S. patent number 5,059,738 [Application Number 07/489,991] was granted by the patent office on 1991-10-22 for method for reactivating mtg process catalyst.
This patent grant is currently assigned to Mobil Oil Corporation. Invention is credited to James H. Beech, Jr., Francis P. Ragonese.
United States Patent |
5,059,738 |
Beech, Jr. , et al. |
October 22, 1991 |
Method for reactivating MTG process catalyst
Abstract
It is disclosed that in a fixed bed process for the conversion
of C.sub.1 -C.sub.4 oxygenates in contact with acidic, medium pore,
shape selective metallosilicate catalyst particles to produce
gasoline boiling range hydrocarbons, including the step of
reactivating spent catalyst at elevated temperature, the cycle time
between regenerations can be substantially improved by reactivating
the spent catalyst at reduced pressure and elevated temperature in
contact with a stream of inert purge gas. Preferably, the
reactivation is carried out at a reduced pressure from
sub-atmospheric to 1400 kPa using nitrogen as a purge gas. Other
purge gases include light paraffinic hydrocarbons, refinery fuel
gas and Group VIII gases of the Periodic Table of the Elements.
Inventors: |
Beech, Jr.; James H.
(Wilmington, DE), Ragonese; Francis P. (Cherry Hill,
NJ) |
Assignee: |
Mobil Oil Corporation (Fairfax,
VA)
|
Family
ID: |
23946143 |
Appl.
No.: |
07/489,991 |
Filed: |
March 7, 1990 |
Current U.S.
Class: |
585/469; 502/56;
585/639; 585/733; 502/34; 585/408; 585/640 |
Current CPC
Class: |
B01J
29/90 (20130101); C10G 3/49 (20130101); C10G
3/62 (20130101); B01J 38/04 (20130101); C10G
3/54 (20130101); C07C 1/20 (20130101); C07C
2529/40 (20130101); C10G 2400/02 (20130101); C10G
2300/703 (20130101); C10G 2300/701 (20130101); Y02P
30/20 (20151101); Y02P 20/584 (20151101); Y02P
20/582 (20151101); C10G 2300/4006 (20130101); C10G
2300/4012 (20130101) |
Current International
Class: |
B01J
38/04 (20060101); B01J 38/00 (20060101); C07C
1/20 (20060101); C07C 1/00 (20060101); C10G
3/00 (20060101); B01J 29/00 (20060101); B01J
29/90 (20060101); C07C 001/20 (); C07C 001/00 ();
B01J 029/38 (); B01J 038/04 () |
Field of
Search: |
;502/34,56
;585/469,639,640,733,408 |
References Cited
[Referenced By]
U.S. Patent Documents
Primary Examiner: Konopka; Paul E.
Attorney, Agent or Firm: McKillop; Alexander J. Speciale;
Charles J. Wise; L. G.
Claims
What is claimed is:
1. In a fixed bed process for the conversion of C.sub.1 -C.sub.4
oxygenates in contact with acidic, medium pore, shape selective
zeolite catalyst particles to produce gasoline boiling range
hydrocarbons, including the step of reactivating spent catalyst at
super atmospheric process pressure and elevated temperature, the
improvement comprising:
reactivating said spent catalyst at pressure less than about 1400
kPa and temperature between about 300.degree. C. and 400.degree. C.
in contact with a stream of inert purge gas for a time period of
from about 4 hours to about 40 hours.
2. The process of claim 1 wherein said improvement comprises a
reactivating pressure of about atmospheric pressure.
3. The process of claim 1 wherein said inert purge gas includes
N.sub.2, light paraffinic hydrocarbons, refinery fuel gas and Group
VIII gases of the Periodic Table of the Elements.
4. The process of claim 1 where in the improvement comprises
reactivating said spent catalyst using predominantly nitrogen purge
gas.
5. The process of claim 1 wherein the improvement comprises
reactivating said spent catalyst at a temperature of about
360.degree. C.
6. The process of claim 1 wherein said catalyst comprises
aluminosilicate having a constraint index of 1-12 and a
silica/alumina molar ratio of at least 12.
7. The process of claim 1 wherein said catalyst comprises
ZSM-5.
8. A process for the conversion of lower aliphatic oxygenated
hydrocarbon to gasoline or gasoline and distillate boiling range
hydrocarbons, comprising;
a) passing said feedstock through a reactor containing a fixed bed
of acidic shape selective medium pore zeolite catalyst particles at
conditions sufficient to effect said conversion;
b) separating effluent from said reactor and recovering said
gasoline or gasoline and distillate boiling range hydrocarbons;
c) interrupting feedstock flow to said reactor upon catalyst
deactivation and loss in feedstock conversion rate;
d) passing inert purge gas to said reactor at temperature between
about 300.degree. C. and 400.degree. C. and pressure less than
about 1400 kPa for a time period of from about 4 hours to about 40
hours whereby deactivated catalyst in said reactor is
reactivated.
9. The process of claim 8 wherein said inert purge gas is passed to
said reactor at about atmospheric pressure.
10. The process of claim 8 wherein said inert purge gas includes
N.sub.2, light paraffinic hydrocarbons and Group VIII gases of the
IUPAC Periodic Table of the Elements.
11. The process of claim 8 wherein said spent catalyst is
reactivated using nitrogen as purge gas.
12. The process of claim 8 wherein said spent catalyst is
reactivated at a temperature of about 360.degree. C.
13. The process of claim 8 wherein said catalyst comprises
aluminosilicate having a constraint index of 1-12 and a
silica/alumina molar ratio of at least 12.
14. The process of claim 8 wherein said catalyst comprises
ZSM-5.
15. The process of claim 8 wherein said feedstock comprises
methanol, dimethylether or mixtures thereof and said hydrocarbons
comprise C.sub.5 + gasoline boiling range hydrocarbons.
16. A process for the reactivation of fixed bed of spent acidic,
medium pore, shape selective zeolite catalyst particles employed in
a process for the catalytic conversion of C.sub.1 -C.sub.4
oxygenates feedstream to gasoline boiling range hydrocarbons,
comprising:
contacting said spent catalyst with a stream of inert purge gas in
the absence of said oxygenates feedstream and at a temperature
between 200.degree. C. and 400.degree. C. and pressure less than
1400 kPa for a time period of from about 4 hours to about 40 hours,
sufficient to reactivate said spent catalyst.
17. The process of claim 16 wherein said reduced pressure is about
atmospheric pressure.
18. The process of claim 16 wherein said inert purge gas includes
N.sub.2, light paraffinic hydrocarbons, refinery fuel gas and Group
VIII gases of the Periodic Table of the Elements.
19. The process of claim 16 wherein said spent catalyst is
reactivated using nitrogen as purge gas.
20. The process of claim 16 wherein said spent catalyst is
reactivated at a temperature of about 360.degree. C.
Description
This invention relates to the process for the zeolite catalyzed
conversion of oxygenated hydrocarbons to higher molecular weight
products. The invention particularly relates to the ZSM-5 catalyzed
fixed bed conversion of methanol or dimethylether to gasoline
boiling range hydrocarbons. More particularly, the invention
relates to an improved process for reactivating or regenerating
spent zeolite catalyst in the methanol-to-gasoline (MTG)
process.
BACKGROUND OF THE INVENTION
Processes for converting lower alcohols such as methanol to
hydrocarbons are known and have become of great interest in recent
times because they offer an attractive way of producing liquid
hydrocarbon fuels, especially gasoline, from sources which are not
of liquid petroliferous origin. In particular, they provide a way
by which methanol can be converted to gasoline boiling range
products in good yields. The methanol, in turn, may be readily
obtained from coal by gasification to synthesis gas and conversion
of the synthesis gas to methanol by well-established industrial
processes. As an alternative, the methanol may be obtained from
natural gas by other conventional processes.
The conversion of methanol to hydrocarbon products may take place
in a fluidized bed process as described, for example, in U.S. Pat.
Nos. 4,071,573 and 4,138,440, or in a fixed bed as described in
U.S. Pat. Nos. 3,998,899, 3,931,349 and 4,035,430. In the fixed bed
process, the methanol is usually first subjected to a dehydrating
step, using a catalyst such as gamma-alumina, to form an
equilibrium mixture of methanol, dimethyl ether (DME) and water.
This mixture is then passed over a catalyst such as zeolite ZSM-5
which brings about the conversion to the hydrocarbon products which
are mainly in the range of light gas to gasoline. The water may be
removed from the methanol dehydration products prior to conversion
to hydrocarbons as may the methanol which can be recycled to the
dehydration step, as described in U.S. Pat. No. 4,035,430. Removal
of the water is desirable because the catalyst may tend to become
deactivated by the presence of the water vapor at the reaction
temperatures employed, but this step is by no means essential.
The conversion of oxygenates, or more particularly methanol, to
gasoline, typically referred to as the MTG process, is highly
energy efficient. The hydrocarbons from the conversion contain 95
percent of the energy in the original methanol feed; the other five
percent is released as exothermic heat and used during the
conversion reaction. Recycling of process gas limits the
temperature rise across the fixed catalyst bed to less than
95.degree. C. Also during the reaction, a small amount of
hydrocarbon is deposited on the catalyst as coke, requiring
periodic catalyst regeneration. Operation of the process, however,
is continuous because additional reactors, arranged in parallel,
permit an individual reactor to swing from operation to
regeneration while another goes from regeneration to operation. The
final gasoline yield from the fixed bed process, after alkylating
the light olefins formed, is about 85-90 percent by weight of the
total hydrocarbons formed. The remaining hydrocarbons are available
mostly as liquid petroleum gas (LPG) and a small amount of fuel
gas.
The conversion of oxygenates is described in depth by C. D. Chang,
Catal. Rev.-Sci. Eng., 25, 1 (1983) and in U.S. Pat. Nos. 3,931,349
to Kuo and 4,404,414 to Penick et al. These references are
incorporated herein in their entirety.
Major problems facing research workers in the field of the MTG
process include improvements in cycle average yield of gasoline and
the extension of catalyst cycle life. Improvements in yield and
catalyst life are known to be inextricably related, whereby
advances in one problem area are typically achieved at the expense
of the other. Process improvements leading to the common
enhancement of gasoline yield and catalyst life have been most
elusive. One factor that complicates the effort of research workers
to achieve the desired advances in yield and catalyst cycle life is
the requirement that the MTG process operate at or near
quantitative methanol conversion. Less than quantitative
conversions, or "methanol breakthrough," presents severe problems
in waste disposal and/or methanol recovery which quickly leads to
punishing economic penalties and, therefore, is to be avoided.
Accordingly, whatever advances research workers are to make in
yield and catalyst life for MTG improvements must be made while
maintaining essentially quantitative conversion of methanol.
The MTG process is a gas phase process which produces isoparaffins
and aromatics with virtually no oligomer formation. The coke formed
in the process is harder and more condensed than that from other
processes such as zeolite catalyzed olefin oligomerization.
Accordingly, it is much harder to remove. It is, however, well
known that deactivated or aged zeolite oxygenate conversion
catalyst can be regenerated by contacting the catalyst at elevated
temperature with an oxygen-containing gas such as air to effect
controlled burning of coke from the deactivated catalyst. While
such a conventional regeneration procedure can restore catalytic
activity diminished by coke formation in the catalyst during the
conversion reaction, regeneration in this manner can lead to
catalyst damage requiring more frequent, and expensive, catalyst
replacement. There is, therefore, a continuing need to find better
methods to regenerate deactivated catalyst in fixed beds in order
to lengthen time on stream, or cycle length.
Accordingly, it is an object of the present invention to provide a
more effective and non-oxidizing process for the reactivation of a
fixed bed containing spent or deactivated zeolite catalyst.
Another object of the invention is to provide a catalyst
reactivation method which would extend the cycle time between
regeneration of fixed bed zeolite catalyst compared to conventional
reactivation methods.
Yet another object of the present invention is to provide a process
more effective and useful for the reactivation of deactivated
zeolite catalyst for lower oxygenate conversion processes.
SUMMARY OF THE INVENTION
The discovery has been made that in a fixed bed process for the
conversion of C.sub.1 -C.sub.4 oxygenates in contact with acidic,
medium pore, shape selective metallosilicate catalyst particles to
produce gasoline boiling range hydrocarbons, including the step of
reactivating spent catalyst at super atmospheric pressure and
elevated temperature, the cycle time between regenerations can be
substantially improved by reactivating the spent catalyst at
reduced pressure and elevated temperature in contact with a stream
of inert purge gas. Preferably, the reactivation is carried out at
a reduced pressure from sub-atmospheric to 1400 KPa using nitrogen
as a purge gas. Other purge gases include light paraffinic
hydrocarbons and Group VIII gases of the IUPAC Periodic Table of
the Elements.
More particularly, the invention comprises a process for the
conversion of lower aliphatic oxygenated hydrocarbon to gasoline or
gasoline and distillate boiling range hydrocarbons. The novel
process includes the steps of passing the feedstock through a
reactor containing a fixed bed of acidic, shape selective, medium
pore zeolite catalyst particles at conditions sufficient to effect
said conversion; separating effluent from the reactor and
recovering gasoline or gasoline and distillate boiling range
hydrocarbons; interrupting the feedstock flow to the reactor
coincident with catalyst deactivation and loss in feedstock
conversion rate; and passing inert purge gas to the reactor at
elevated temperature and reduced pressure whereby deactivated
catalyst in the reactor is reactivated.
In the process of the present invention, spent or otherwise
deactivated catalyst is reactivated at a temperature between
300.degree. C. and 400.degree. C. at about atmospheric pressure
using nitrogen as purge gas. The preferred reactivation temperature
is 360.degree. C.
DESCRIPTION OF THE FIGURES
FIG. 1 is a process flow diagram of the MTG process.
FIG. 2 is a plot of reactor midpoint temperature positions showing
the recovery of activity for Examples 1 and 2 of the instant
invention.
FIG. 3 is a plot of reactor midpoint temperature position for the
standard reactivation procedure.
FIG. 4 is a plot of methanol conversion history for Example 1 of
the invention.
FIG. 5 is a plot of methanol conversion history for Example 2 of
the invention.
DETAIL DESCRIPTION OF THE INVENTION
The MTG process is useful for the conversion of a number of
differing oxygenated organic compounds into hydrocarbon products.
The process is useful for the conversion of aliphatic compounds
including lower alcohols such as methanol, ethanol and propanol,
ethers such as DME and diethyl ether, ketones such as acetone and
methyl ethyl ketone, aldehydes such as acetaldehyde, esters such as
methyl formate, methyl acetate and ethyl acetate, carboxylic acids
such as acetic acid, butyric acid and their anhydrides e.g., acetic
anhydride. Examples of conversions of such compounds may be found
in U.S. Pat. Nos. 3,907,915, 3,894,107, 3,894,106, 3,894,103,
3,894,104, and 3,894,105 to which reference is made for details of
the conversions. The product in each case will be a hydrocarbon
mixture ranging from light gas to heavier fractions (C.sub.10+) but
will generally be concentrated in the gasoline boiling range
(C.sub.5 -220.degree. C.). The process is particularly useful in
the catalytic conversion of methanol to hydrocarbons in the
gasoline boiling range and, for convenience, the process will be
described below with reference to such a process although it should
be remembered that the principles are applicable to a broader range
of conversion, as set out above.
If methanol is used as the starting material for the process it is
preferred to subject it to an initial dehydration step to form an
intermediate product comprising dimethyl ether (DME). The DME is
then passed to the hydrocarbon step with either complete, partial
or no separation of the unreacted methanol and the water produced
as a by-product of the dehydration. However, it is not essential to
carry out this dehydration even though it is preferred. It is
possible to dehydrate only part of the methanol with, for example,
the dehydration product going to one reactor and the raw methanol
going to another.
Because the oxygenated charge may be fed into the reactors in
different forms, e.g., methanol and DME, it will often be
convenient, for purposes of calculating recycle ratio and other
factors, to base the calculations upon a single equivalent charge.
For example, if both methanol and DME are fed to the reactors, the
total charge may be reduced to a basis of "methanol equivalents" in
which one mole of DME is equal to two methanol equivalents. Thus,
the reactant flow at any point may be readily reduced to a single
value from which others may be derived, e.g., recycle ratio, or
reactant or feedstock feedrate expressed as weight hourly space
velocity (WHSV) based on catalyst.
The conversion of methanol or methanol equivalents to gasoline is
accomplished in contact with zeolite catalysts, such as ZSM-5,
usually quantitatively in the presence of active catalyst and, in
the process of the present invention, in a fixed catalyst bed. In
addition to gasoline and other hydrocarbons, water is a reaction
by-product. However, process variables must be carefully managed
because the conversion of methanol to gasoline boiling components
is a highly exothermic reaction releasing approximately 750 BTU of
heat per pound of methanol. This amount of heat release will result
in an adiabatic temperature increase of about 1200 degrees F. for
pure methanol feed. In an adiabatic catalyst bed reactor, this
large temperature increase will result in high catalyst aging
rates, and possibly cause thermal damage to the catalyst.
Furthermore, such high temperatures could cause an undesirable
product distribution to be obtained. Therefore, it is critical to
the conversion of methanol to useful products to provide sufficient
heat removing or dissipating facilities particularly during initial
contact with the crystalline zeolite conversion catalyst so that
the maximum temperature encountered in any portion of the zeolite
catalyst conversion step is below an upper predetermined limit.
The exothermic character of the conversion reaction also requires
careful management of the methanol feedrate in terms of weight
hourly space velocity (WHSV) based on catalyst loading. "Methanol
breakthrough," a term of art indicating the appearance of methanol
in the aqueous product stream and, therefore, less than
quantitative conversion, has generally been followed to signal the
end of the process cycle and the need to regenerate catalyst. The
production of even very dilute aqueous methanol product streams
presents an operator with very costly waste disposal or separation
problems and must be avoided. Accordingly, when any combination of
process parameters produce a methanol cycle under those conditions
is ended largely for economic reasons. Of course, if disposal or
recovery of unconverted methanol is not a consideration, the cycle
can be continued to less than 99% methanol conversion.
Referring now to FIG. 1 a typical process flow diagram of the MTG
process is presented. Crude methanol in a liquid phase condition is
charged to the process by conduit 2 communicating with pump 4. The
methanol is pressured to about 1500-5000 kPA, preferably 2500 kPa
(350 psig), in pump 4 and then passed by conduit 6 to heat
exchanger 8 wherein the liquid methanol is preheated. It is then
passed into drum 9 where it is vaporized at about 185.degree. C.
(400.degree. F.) by indirect heat exchanger 11. The methanol is
then superheated in indirect exchanger 13.degree. to about
315.degree. C. (600.degree. F.) and it is passed by conduit 10 to
the inlet of the dimethyl ether forming catalytic reactor 12. In
catalyst contained in reactor 12, a fixed bed of gamma alumina
catalyst is maintained as a fixed bed of catalyst through which the
methanol reactant passed downwardly through or as an annular bed of
catalyst for radial flow of reactant material therethrough. A
single down-flow fixed catalyst bed or a plurality of separate
fixed downflow catalyst beds are arranged for converting the
methanol feed under restricted temperature conditions as herein
described to essentially an equilibrium product comprising
methanol, dimethyl ether or water at a temperature of about
395.degree.-415 .degree. C. (740.degree.-780.degree. F.) due to the
exothermic temperature rise catalytically generated in the
operation. The equilibrium product thus obtained may be construed
as an ether rich product which is then passed by conduit 14 to a
second reactor stage 16 housing one or more separate and
sequentially arranged beds of a ZSM-5 type of crystalline
zeolite.
A diluent material introduced by conduit 18 is combined with the
ether rich effluent obtained as hereinbefore discussed before
contact of the mixture is made with the HZSM-5 crystalline zeolite
catalyst under heat generating or exothermic reaction conditions
controlled to restrict the temperature increase between the reactor
inlet and reactor outlet not to exceed about 94.degree. C.
(200.degree. F.) and preferably not to exceed about 65.degree. C.
(150.degree. F.). The conversion of the ether rich effluent by the
HZSM-5 catalyst is highly exothermic as discussed above and
controlled within desired limits by use of gasiform heat
dissipating diluent material. During this highly exothermic
operation the ether rich effluent or equilibrium mixture comprising
dimethyl ether, methanol and water is controlled to effect the
conversion thereof to gasoline boiling range components comprising
aromatic and isoparaffins. The aromatic components comprising
benzene, toluene and xylene are preferred components over the
higher boiling durene aromatic material and efforts are made
(e.g.,reactant partial pressure, space velocity and reactant plug
flow operation) to promote this end.
The product effluent of the HZSM-5 reaction zone 16 is passed
through one or more cooling steps to reduce the temperature to a
desired low temperature. In the specific arrangement of the figure
the effluent is passed by conduit 20 to heat exchanger 22 wherein
the effluent temperature is reduced to about 94.degree. C.
(200.degree. F.) by indirect heat exchange with diluent material
removed therefrom by conduit 18. The diluent will be at a
temperature of about 315.degree.-343.degree. C.
(600.degree.-650.degree. F.). The partially cooled effluent is
removed from heat exchanger 22 and passed by conduit 24 to cooling
water and/or air heat exchanger 26 wherein a further cooling of the
effluent to about 38.degree. C. (100.degree. F.) is accomplished.
Some of the effluent is passed via conduit 21 to heat exchangers
13, 11, and 8 to superheat, vaporize, and preheat, respectively,
the methanol feed. The effluent from exchanger 8 is cooled in
exchanger 26 and combined with cooled effluent from reactor conduit
20 and passed into separator 28, where liquid hydrocarbon, liquid
water and gaseous material are separated. In the arrangement of the
drawing, most of the gaseous effluent is then passed by conduit 30
to heat exchanger 22 where it is again passed in indirect heat
exchange with reactor effluent and finally heater 38 before
entering reactor 16. Water product is removed from separator 28 via
conduit 34 for further treatment. Liquid hydrocarbon product is
removed from separator 28 via conduit 32 and is sent to a product
recovery section (not shown). Of course many other heat exchange
arrangements may be provided for reducing the reactor effluent
temperature from about 426.degree. C. (800.degree. F.) to about
38.degree. C. (100 degrees F.) before passage to separator 28.
Separator 28 is maintained at a temperature of about 38.degree. C.
(100.degree. F.) and a pressure of about 1540 kPa (220 psig). In
the separator a rough cut is made between gasiform diluent
materials, desired aromatic and isoparaffin product and water.
Water is withdrawn by conduit 34. A gasiform product material lower
boiling than desired gasoline boiling range constituents is
withdrawn by conduit 30 and passed to a compressor 36. A plurality
of parallel arranged gas compressors may be used for this purpose.
The gasiform material is compressed by compressor 36 to a pressure
of about 2310 kPa (330 psig) before passage to exchanger 22. Excess
gas is removed via conduit 39 and sent to product recovery.
The conversion of methanol or methanol equivalents is preferably
catalyzed by a crystalline zeolite catalyst having acidic
functionality. The preferred class of catalysts is characterized by
a Constraint Index of 1 to 12 and a silica:alumina ratio of at
least 12:1 and preferably higher e.g. 20:1 to 200:1 or even higher.
As described in U.S. Pat. No. 3,998,889, the Constraint Index of a
zeolite is a convenient measure of the extent to which a zeolite
provides constrained access to its internal structure for molecules
of different sizes. It is therefore a characteristic of the
structure of the zeolite but is measured by a test which relies
upon the possession of cracking activity by the zeolite. The sample
of zeolite selected for determination of the Constraint Index of a
zeolite should therefore represent the structure of the zeolite
(manifested by its X-ray diffraction pattern) and have adequate
cracking activity for the Index to be determined. If the cracking
activity of the selected zeolite is too low, the Constraint Index
may be determined by using a zeolite sample of the same structure
but higher cracking activity which may be obtained, for example, by
using an aluminosilicate zeolite of higher aluminum content.
Details of the method of determining Constraint Index and of the
values of the Index for typical zeolites are given in U.S. Pat. No.
3,998,899 to which reference is made for such details and other
information in this respect.
The silica-alumina ratios referred to in this specification are the
structural or framework ratios, that is, the ratio for the
SiO.sub.4 to the AlO.sub.4 tetrahedra which together constitute the
structure of which the zeolite is composed. This ratio may vary
from the silica:alumina ratio determined by various physical and
chemical methods. For example, a gross chemical analysis may
include aluminum which is present in the form of cations associated
with the acidic sites on the zeolite, thereby giving a low
silica:alumina ratio. Similarly, if the ratio is determined by
thermogravimetric analysis (TGA) of ammonia desorption, a low
ammonia titration may be obtained if cationic aluminum prevents
exchange of the ammonium ions onto the acidic sites. These
disparities are particularly troublesome when certain treatments
such as dealuminization methods which result in the presence of
ionic aluminum free of the zeolite structure are employed to make
highly siliceous zeolites. Due care should therefore be taken to
ensure that the framework silica: alumina ratio is correctly
determined.
Preferred zeolites which have the specified values of Constraint
Index and silica:alumina ratio include zeolites ZSM-5, ZSM-11,
ZSM-12, ZSM-35, and ZSM-48, which are described in U.S. Pat. Nos.
3,702,886 (ZSM-5), 3,709,979 (ZSM-11), 3,832,449 (ZSM-12),
4,076,842 (ZSM-23) and 4,016,245 (ZSM-35), and European Patent
Publication No. 15132, and reference is made to these patents for
details of these zeolites, their preparation and properties. Of
these zeolites, ZSM-5 is preferred.
The zeolite catalyst used is at least partly in the hydrogen form
e.g. HZSM-5 but other cations e.g. rare earth cations may also be
present. When the zeolites are prepared in the presence of organic
cations they may be quite inactive possibly because the
intracrystalline free space is occupied by the organic cations from
the forming solution. The zeolite may be activated by heating in an
inert atmosphere to remove the organic cations e.g. by heating at
over 500 degrees C. for 1 hour or more. The hydrogen form can then
be obtained by base exchange with ammonium salts followed by
calcination e.g. at 500 degrees C. in air. Other cations e.g. metal
cations can be introduced by conventional base exchange
techniques.
In the conventional MTG process as practiced in the art heretofore,
cycle average gasoline yields generally run about 80 to 85%.
Typical cycle lengths are between 20 to 40 days before methanol
breakthrough at 99.9% conversion occurs. Of course, virtually every
process parameter can effect yield and cycle life, at least
negatively, but by and large, catalyst deactivation has a
dominating effect on these properties of the process. During normal
operation of the MTG reactors the zeolite catalyst undergoes
gradual aging and deactivation, associated with the deposition of
coke and carbonaceous materials on the catalyst. In the fixed bed
system catalyst deactivation by coke causes a movement of the
reaction zone toward the reactor outlet, an effect normally
referred to as "band aging". This movement can be tracked by
recording the catalyst bed position at which one-half the
temperature rise has occurred, i.e., the temperature profile
midpoint. A decrease in the value of the temperature profile
midpoint is indicative of an improvement of catalyst activity
during a cycle. Reactivation of catalyst will result in movement of
the reactor temperature profile midpoint toward the reactor
inlet.
Spent or deactivated catalyst is reactivated by swinging the
reactor from process to regeneration mode by terminating the flow
of feedstock to the fixed bed reactor and regenerating the catalyst
by contacting it with air or other oxygen-containing gas at high
temperature. As an alternate method of reactivation, it is known
that a recycle gas flow, i.e., C.sub.4 - hydrocarbon and hydrogen ,
can be passed in contact with deactivated catalyst at operating
conditions (2100 kPa and 350.degree. C.) for about 24 hours. Under
these conditions, catalyst activity, as measured by reactor
temperature midpoint, recovers from 60% midpoint position, or
closer to reactor outlet, to a 45% midpoint position, or close to
reactor inlet, after deactivation and reactivation in a commercial
plant operation. The interrupted flow of feedstock to the reactor
is reinitiated after reactivation. Accordingly, the multiple fixed
bed reactors in the process alternates between catalyst
regeneration and process conversion conditions.
When used herein in relation to a zeolite which has undergone aging
in an oxygenate conversion process, the term `reactivate` is
intended to mean that, after reactivation, there is an increase in
the rate at which the zeolite converts oxygenates to gasoline
boiling-range hydrocarbons.
The aforementioned methods of catalyst reactivation in the MTG
process have distinct disadvantages. Oxidative regeneration of
spent catalyst can result in catalyst damage due to the high
temperatures and water formation in the combustion process. While
regeneration with a recycle gas at operating conditions obviates
some of the problems associated with oxidative regeneration, the
extent of coke removal is less.
The new and improved reactivation technique of the present
invention results in a greater activity recovery by conducting
reactivation at low pressure, i.e., atmospheric pressure, with an
inert gas purge rather than at elevated pressure with a recycle
hydrocarbon gas purge. An improvement in regeneration under the
high temperature and low pressure regeneration conditions of the
present invention represents an unexpected discovery.
The following Examples 1 and 2 are provided to illustrate the
present invention compared to standard reactivation procedures
illustrated in Example 3.
EXAMPLE 1
An MTG process reactor, operating under typical conversion
conditions of 680.degree. F. and 2100 KPa is shut down and the flow
of methanol to the process interrupted. The catalyst is allowed to
cool slowly over six to ten hours while the pressure is at
atmospheric under a flow of nitrogen purge gas. The reactor is
brought back on stream by initiating flow of methanol to the
reactor under the aforementioned conversion conditions. When
finally brought back on stream the midpoint of the temperature
profile position had changed from 33% before shut down to 18%. This
indicates that under the flow of inert gas at reduced pressure
catalyst had been reactivated to move the midpoint of the
temperature profile closer to the reactor inlet. At the end of the
cycle the procedure was repeated, shown in Example 2, and the
reactivation occurred again with the midpoint position changing
from 42% to 28%. Methanol conversion also responded favorably to
the reactivations by increasing.
The following Table 1 presents the conditions and results for
Examples 1, 2 and 3 of the present invention for catalyst
reactivation and compares them with the conditions in Example
3.
TABLE 1 ______________________________________ Example 1 2 3
______________________________________ Midpoint, % Before React. 33
42 63 After React. 18 28 53 Relative % change 45 33 16 Methanol
Conversion Before, wt % 99.98 99.78 99.86 After, 99.99 99.89 99.94
Change, +0.01 +0.11 +0.08 Methanol in Water Product Before, ppm wt.
210 2790 1840 After, 160 1400 800 Days on stream 26 42 52 at
Reactivation Reactivation Temp., .degree.F. 680 to 80 680 660
Press., PSIG 0 0 300 Gas Type N2 hydrocarbon Duration, hrs 10 16 24
______________________________________
The results from Examples 1, 2 and 3 are presented graphically in
FIGS. 2, 3, 4, and 5. Reactor temperature midpoint positions,
showing the recovery of activity, are plotted in FIG. 2 for
Examples 1 and 2. The case for the standard reactivation procedure,
Example 3, is plotted in FIG. 3. FIGS. 4 and 5 show the methanol
conversion increase after reactivation for Examples 1 and 2 that
illustrate the process of the present invention.
Comparison of FIG. 3 with FIG. 2 clearly shows that the process of
the present invention as plotted for Examples 1 and 2 in FIG. 2
results in a substantially greater change in temperature profile
midpoint of approximately 45% and 33%, where the case for Example 3
shows a midpoint position change of approximately 16%. Referring to
Table 1, the new activation procedure is shown to result in an
increase methanol conversion for Example 1, +0.11% for Example 2,
and 0.08 for Example 3.
The reactivation process of this invention can be carried out using
an inert gas purge comprising nitrogen, paraffinic hydrocarbons,
refinery fuel gas, carbon dioxide, or gases of Group VIII of the
Periodic Table of the Elements. Nitrogen (N.sub.2) is the preferred
inert purge gas. The process can be carried out at pressures from
subatmospheric to about 400 KPa, but operation at atmospheric
pressure is preferred. Reactivation can be accomplished at
conventional operating temperatures between 200.degree.-400.degree.
C. Preferably, reactivation is initiated at approximately regular
operating temperature of 360.degree. C. The reactivation is carried
out over a period of from 4-40 hours, but preferably between about
10 and 16 hours.
This new reactivation technique increases the magnitude of the
benefits derived from the standard reactivation technique by moving
the reactive catalyst zone (temperature midpoint) closer to the
inlet of the reactor and increasing methanol conversion. This new
method may allow for multiple reactivations before an oxygen
regeneration is needed. This possibility is enhanced, for example,
by the availability of vacuum facilities in many commercial
regeneration systems. This will result in reduced operating costs
and increased catalyst life which will reduce catalyst costs. Since
no thermal shocking of the catalyst is involved because temperature
is kept constant (unlike during oxygen regenerations) catalyst
breakage should be reduced.
While the instant invention has been described by specific examples
and embodiments, there is no intent to limit the inventive concept
except as set forth in the following claims.
* * * * *